Selective conversion process



3,395,096 SELECTIVE CGNVERSION PROCESS Elroy M. Gladrow, Baton Rouge, Ralph Burgess Mason, Denham Springs, and Glen Porter Hamner, Baton Rouge, La., assignors to Esso Research and Engineering Company, a corporation of Delaware No Drawing. Filed June 7, 1966, Ser. No. 555,716 20 Claims. (Cl. 208111) ABSTRACT OF THE DISCLOSURE Straight-chain hydrocarbons selectively hydrocracked with rare earth metal-containing crystalline zeolite having pore openings of less than 6 Angstrom units. Catalyst can also contain hydrogen and/or Group II-B metal cations. Useful for improving octane rating of naphtha fractions and dewaxing middle distillate fractions.

This invention relates to the removal of straight-chain hydrocarbons from petroleum-derived feedstocks by their selective conversion in the presence of hydrogen. More particularly, it relates to a selective hydroc'racking process which is accomplished-in the presence of a rare earth metalcontaining crystalline metallo .alumino-silicate having uniform pore openings less than about 6 Angstrom units in diameter, preferably about 5 Angstroms.

I Hydrocarbon conversion and upgrading with crystalline alumino-silicate zeolite catalysts are now well known in the art. The use of these materials for such purposes as hydrocracking has been generally directed to typical petroleum-derived feedstocks such as gas oils, etc., which are customarily converted to lower boiling productsuseful as gasoline. The crystalline zeolites employed for such purposes usually have uniform pore openings of about 6 to Angstroms and are therefore non-selective; that is, substantially all of the feed molecules are admitted into the zeolite pore structure and converted. For many purposes selective hydrocracking. of particular molecular species in the feed to the substantial exclusion of others is desired. For example, octane improvement of naphtha fractions can be accomplished by selectively hydrocracking only the straight-chain hydrocarbons (e.g., olefins, paraflins, etc.) which tend to be low octane producing, thereafter removing the hydrocracked products, and finally recovering a higher octane product. Selective hydrocracking of straight-chain hydrocarbons contained in lube oil or gas oil fractions is also valuable for the purpose of pour point reduction or dewaxing. The use of a non-selective large pore (e.g., 6 to 15 Angstroms) crystalline zeolite for such purposes is largely ineffectual, as the desired feed molecules (e.g., aromatics) are admitted into the zeolite pores and converted together with the straight-chain hydrocarbons.

With specific reference to the upgrading of naphtha fractions for inclusion in the high quality motor gasoline necessary for modern automobiles, it is customary to improve the octane rating and cleanliness or gum-forming properties by means of such processes as thermal or catalytic reforming. The degree of octane improvement by reforming is usually limited by formation of coke and gas as reaction temperature increases. Similarly, octane improvement of olefinic naphthas by other means is often unavailing; e.g., catalytic cracking results in an undesirable high gas and coke make, and hydrofining results in an octane number loss. Attempts at solving these problems have generally involved one ormore hydro techniques, such as hydrocracking, hydroforming, hydrodealkylation, etc., which processes tend to form lesser amounts of coke and dry gas while at the same time resulting in improved octane product. However, in-

States Patent 0 ice discriminate use of hydrocracking, for example, is often self-defeating, since products boiling below the naphtha range are formed, and naphtha yield is thereby reduced. Hydroforming or catalytic reforming is also not practical with certain naphtha feeds, e.g., coker naphthas, which contain appreciable sulfur, nitrogen and diolefins, again because of excessive coke make and rapid catalyst deactivation. Catalytic hydroforming, which depends upon aromatics formation for octane improvement, is ineffective with feeds having low cycloparaffin concentration.

In view of the above problems, it will be realized that a selective conversion process capable of removing low octane-producing components in the naphtha feed, with minimum conversion of high octane-producing compo: nents, is greatly to be desired. Removal of the low octane components would thus result in enhancement of the naphtha octane number without appreciable alteration in boiling range.

With respect to the dewaxing of waxy feeds such as lube stocks, middle distillates and the like for pour point and/ or cloud point reduction, various means are presently available including extraction, adsorption, etc. In the case of adsorption with molecular sieves, certain difficulties are encountered in large scale removal of normal parafiins from branched-chain and cyclic hydrocarbons. For example, it is usually necessary to employ a two-step cyclic process wherein the normal parafi'ins are first selectively adsorbed and then desorbed in a separate operation. Such cyclic processes are relatively expensive because of the frequency with which the sieve must be desorbed. Further, the desorption methods presently available a-re only partially effective as selectivity and capacity rapidly decline with use as carbonaceous deposits accumulate on the adsorbent. Frequent regeneration is required.

The present process for removal of straight-chain hydrocarbons involves the use of certain types of crystalline alumino-silicate zeolites having uniform pore openings of less than 6, preferably about 5, Angstrom units. The precise nature of the zeolites used 'herein will be hereinafter described in detail. The distinctions between the present process and those of the prior art should first be noted. The present process differs from prior adsorption methods in that the molecular sieves are employed to effect chemical conversion of the normal paraffins on a selective basis, as opposed to a mechanical separation. This selective conversion is accomplished in the presence of hydrogen at critical conditions of temperature, pres sure, feed rate and hydrogen rate. The present process has numerous advantages over prior processes which have been proposed for dewaxing hydrocarbon oils. Normal paraflins which would otherwise be adsorbed by the molecular sieve material are continuously converted to lower boiling gaseous products which are not retained within the zeolite pores. Desorption is thus unnecessary. Further, lower boiling products such as kerosene, naphtha, butane, etc. can be readily separated from the normally liquid portions of the dewaxed efliuent and recovered as valuable by-products. The economics of the simplified procedure employed herein are considerably more attractive than previously attainable.

With respect to the use of crystalline zeolites as hydrocarbon conversion catalysts, an extensive body of prior art has collected in recent years. For example, US. Patent Nos. 2,971,903 and 2,971,904 disclose various hydrocarbon conversion processes employing crystalline alumino-silicates having uniform pore openings between about 6 and 15 Angstroms. The zeolites used in the present process have a pore size less than 6 Angstroms, preferably about 5 Angstroms, which pore size has been found to be necessary and critical to the successful selective hydrocracking herein contemplated. The prior art has also recognized the possibility of selectively cracking normal parafiins by means of 5Angstrom molecular sieves for such purposes as dewaxing, etc. These uses derive from the ability of these crystalline zeolite materials to selectively admit certain size molecules while rejecting others. For example, U.S. Patent No. 3,039,953 discloses the selective conversion of normal paraffins with a S-Angstrom zeolite, and US. Patent No. 3,140,322 relates generally to selective catalytic conversion utilizing crystalline zeolites disclosing such processes as dehydration, catalytic cracking and hydrogenation.

The essence of the present invention, which distinguishes it from the above prior art teachings, resides in the surprising discovery that certain unique S-Angstrom crystalline alumino-silicates are superior catalyst components for selective hydroconversion reactions in general and selective hydrocracking in particular. The present invention represents further an improvement over those disclosed and claimed in copending applications Serial Nos. 444,796 and 518,680. Those applications are directed to the discovery of the unexpected superiority of certain hydrogen and/ or Group II-B metal-containing S-Angstrom crystalline zeolites for the removal of straightchain hydrocarbons by selective hydroconversion. Such zeolites, when combined with a suitable hydrogenation component, were previously found to be excellent catalysts for this purpose.

It has now been found that the S-Angstrom type of crystalline zeolite may be still further improved with respect to its selectivity for the conversion of straight-chain hydrocarbons in the presence of hydrogen by the incorporation of rare earth metals, preferably via base exchange. The resulting catalyst has been found to be more effective than the superior forms previously disclosed in the aforementioned copending applications. As a result of this increased effectiveness for this particular purpose, lower operating temperatures can be employed so as to avoid excessive conversion of other desired molecular species in the feed stream. As a result, increased yield of desired product is attainable with maximum removal of straight-chain hydrocarbons by their selective conversion to lower boiling, readily removable materials.

The catalysts utilized in the present invention will now be described in detail. The starting material will be a crystalline alumino-silicate having relatively small uniform pore openings, i.e., less than 6 Angstroms, particularly 4 to less than 6 Angstroms, e.g., about 5 Angstroms. More particularly, the zeolites employed will have uniform pore openings capable of affording entry to the objectionable normal parafiinic hydrocarbons but substantially incapable of admitting the more valuable branched and cyclic hydrocarbons. Preferred crystalline alumino-silicate zeolites in the present invention will include Zeolite A and the natural or synthetic form of erionite. Zeolite A is described in US. Patent No. 2,882,243 as having a molar formula in the dehydrated form of wherein M is a metal usually sodium and n is its valence. It may be prepared by heating a mixture containing N320, A1 SiO and H 0 (supplied by suitable source materials) at a temperature of about 100 C. for 15 minutes to 90 hours or longer. Suitable ratios of these reactants are described in the aforementioned patent. The products will have uniform pore openings of about 4 Angstroms in the sodium form. They may then be converted to materials having uniform pore openings of about 5 Angstroms by replacement of sodium via conventional ion-exchange techniques with various cations. In addition to Zeolite A, other relatively small size zeolites can also be employed, such as the naturallyoccurring mineral erionite which has elliptical pore openings of about 4.7 to 5.2 Angstroms on its major axis. The synthetic form of erionite is also suitable and can be prepared by known methods such as those disclosed in US. Patent No. 2,950,952. Synthetic erionite is characterized by pore openings of approximately 5 Angstrom units and differs from the naturally-occurring form in its potassium content and the absence of extraneous metals. Other natural zeolites having effective pore diameters less than 6 Angstroms, preferably 5 Angstroms, are also contemplated herein, such as chabazite, analcite, mordenite, lebrynite, natrolite, etc. Thus, both the natural and synthetic varieties of S-Angstrom zeolites can be used with the only limitation being one of pore size. As indicated, the pore size must be sufiicient to substantially admit the straight-chain hydrocarbons but insufficient to substantially admit the valuable high octane-producing components, such as the aromatics, in order to avoid their hydrocracking. This capacity should, therefore, be demonstrated at the particular hydrocracking conditions to be employed, as the effective pore diameter of these zeolite materials often varies with temperature and pres sure.

In accordance with the invention, the alkali metal form of the relatively small pore size crystallizine zeolite as it is synthesized, e.g., the sodium and/or potassium form, is treated to incorporate rare earth metal cations. This is conveniently done by base-exchange procedures which are Well known in the art. In the most preferred embodiment, the relatively small pore size zeolite will first be base exchanged with a hydrogen-containing cation, such as ammonium ion, by treatment with a solution of an ammonium salt, such as the chloride, nitrate, sulfate, etc., or via a mild acid treatment. In the case of ammonium ion exchange, subsequent calcination will result in liberation of ammonia and the formation of the hydrogen form of the zeolite. This initial ammonium-exchange step is desired in order to substantially replace the alkali metal ions normally contained within the zeolite structure, due to the greater ease of exchangeability of rare earth cations with hydrogen and/0r ammonium cations, as opposed to alkali metal cations. The extent of the original ammonium ion exchange should be sufiicient to reduce the alkali metal content of the zeolite to less than about 5 wt. percent, preferably less than about 4 wt. percent, so that about 60%, preferably about or more of the alkali metal cations originally contained in the zeolite are replaced with hydrogen-containing cations. Following the hydrogen-containing ion-exchange step, the zeolite is contacted with a suitable rare earth metal compound in order to introduce rare earth metal cations into the zeolite structure by exchange with the hydrogen-containing and/or residual alkali metal cations. It will thus be preferable to utilize compounds wherein the rare earth metal ion is in the cationic state.

A wide variety of rare earth compounds can be employed with facility as a source of rare earth ions. Operable compounds include rare earth chlorides, bromides, iodides, sulfates, thiocyanates, peroxysulfates, acetates, benzoates, citrates, fluorides, nitrates, formates, propionates, butyrates, valerates, lactates, tartrates and the like. The only limitation on the particular 'rare earth metal salt or salts employed is that it be sutliciently soluble in the fluid medium in which it is used to give the necessary rare earth ion transfer. Representative of the rare earth metals are those having atomic numbers from 57 to 71, inclusive, and scandium and yttrium.

The rare earth metal salts employed can either be the salts of a single rare earth metal or mixtures of rare earth metals, such as the rare earth chlorides or didymium chloride. A rare earth chloride solution is a mixture of rare earth chlorides consisting essentially of the chlorides of lanthanum, cerium, neodymium and praseodymium with minor amounts of Samarium, gadolinium and yttrium. Rare earth chloride solutions are commercially available and typically contain the chlorides of the rare earth mixture having the relative composition cerium (as CeO 48% by weight, lanthanum (as Lagog) 24% by weight praseodymium (as Pr O 5% by weight, neodymium (as Md O 17% by weight, samarium (as Sm O 3% by weight, gadolinium (as Gd O 2% by weight and other rare earth oxides 0.8% by weight. Didymium chloride, which is specifically referred to in the examples, is also a mixture of rare earth chlorides but having a lower cerium content. It consists of the following approximate amounts by weight of rare earths determined as oxides: lanthanum 45-46%, cerium l2%, prarseodymium 910%, neodymium 32-33 sarnarium 56%, gadolinium 3-4%, yttrium 0.4% and 1-2% of other rare earths. It is to be understood that other mixtures of rare earths are also applicable for the preparation of the novel compositions of this invention, although lanthanum, neodymium, praseodymium, samarium and gadolinium, as well as mixture of rare earth cations containing a predominant amount of one or more of the above cations, are preferred. The exchange medium utilized ordinarily will be water, although other solvents can be used, assuming that the rare earth metal compound will ionize in that solvent. The concentration of the rare earth metal compound employed in the base-exchange solution will vary depending upon the particular alkali and/ or hydrogen content of the zeolite and on the conditions under which treatment is eifected. However, the ion-exchange procedure should be conducted so as to reduce the alkali metal content of the original zeolite to less thanabout 5, preferably less than about 4 wt. percent and to incorporate about 0.3-10.0, preferably about 1.0 to about 7.5 wt. percent rare earth metal into the zeolite. The rare earth metal cation exchange is preferably conducted under conditions of buffered acidity in the pH nangeof about 3 to 7.0. The temperature at which the base-exchange procedures are effected may vary widely, ranging from room temperature to elevated temperatures below the boiling point of the treating solution. Generally an excess of base-exchange solution will be employed, and the time that the zeolite contacts the base-exchange solution will be governed by the aforementioned ranges. In addition to the sequential base-exchange treatment involving initial treatment with hydrogen-containing cations, followed by rare earth cation treatment, base-exchange solutions can be employed which contain both the hydrogen-containing cations and the rare earth metal cations, so long as the [required ranges are satisfied.

In addition to the hydrogen-rare earth form, it is'further contemplated herein that the rare earth metal-containing zeolite also contain a Group IIB metal cation. In certain cases and/or with particular feedstocks, this form of the zeolite may be desired. For example, when the alkali metal cations have been largely removed by the hydrogen-containing ions, complete replacement by rare earth ions may not be economical. Further, the incorporation of Group IIB cations will often provide for greater selectivity in the conversion of naphthas, probably from the slightly smaller pore diameter which results. In this instance the rare earth metal exchange can be accomplished with the alkali metal zeolite or with the hydrogen-containing form zeolite, followed by suitable exchange with a Group II-B metal cation solution. Suitable Group II-B metals include cadmium and zinc cations, with zinc cations being the more preferred. For example, a preferred cation solution will be an aqueous solution of a zinc salt, such as zinc chloride, zinc acetate, etc. In this embodiment the extent of ion exchange should be sufficient to result in a final zeolite product wherein the alkali metal content has been reduced to less than about 5.0 wt. percent, preferably less than 2.5 wt. percent, with the remainder of the cationic content of the zeolite being composed of the rare earth and the Group IIB metal ions, with optional amounts of hydrogen. The preferred weight ratio of Group II-B metal to rare earth metal will be in the range of about 0.5:1 to about 10:1, more preferably about 1:1 to about 7:1.

In generaL'therefor'e, the end result of the various ionexchange steps should be to reduce the alkali metal content to the aforementioned'ranges, and the ion exchange will preferably be conducted to cause at least 60%, preferably greater than of the exchangeable cation content to be comprised of the rare: earth and hydrogencontaining and/or Group II-B metal cations. To summarize, the selective hydrocracking of straight-chain hydrocarbons in accordance with the present invention is accomplished with a catalyst comprising a relatively small pore size crystalline zeolite which has been base exchanged with both hydrogen-containing cations and rare earth metal cations, or with Group II-B metal cations and rare earth metal cations, or with hydrogen-containing cations, rare earth metal cations and Group II-B metal cations, or with rare earth metal cations alone.

As a further step in the preparation of the zeolitic materials used in the present process, the exchanged zeolite is preferably combined with an active hydrogenation metal component chosen from Groups V B, VI-B, VIIB or VIII of the Periodic Table. Such hydrogenation components are suitably exemplified by the metals cobalt, nickel, platinum, palladium, etc. These metals may exist in the form of the free metal, or the oxide or sulfide, or mixtures of such metals, oxides or sulfides. Platinum group metals (i.e., metals of the platinum and palladium series) will be preferred in the present invention, with palladium being particularly preferred. Incorporation of the active metal may be accomplished by any conventional technique, such as by ion exchange followed by reduction, impregnation, etc. When palladium is employed, the zeolite is preferably contacted with an ammoniacal solution of palladium chloride (Pd(NH Cl in sufficient amount to produce the desired amount of palladium in the final product, and then dried at low temperature and calcined at a temperature of 800 to 1000 F. Thus, where the zeolite has been previously exchanged with ammonium ion, calcination at this point serves to liberate ammonia and produce the hydrogen form of the zeolite. In the case of platinum group metals, a reduction step will usually be preferred, for example, by treatment with hydrogen. The amount of the active hydrogenation metal component may range from about 0.1 to about 25 wt. percent, based on the weight of the final product. In the case of platinum group metals, e.g., palladium, the preferred amount will be in the range of about 0.1 to 6, e.g., 0.3 to 1.3, wt. percent, based on dry zeolite. While the presence of the added hydrogenation component is preferred, it will not always be necessary, particularly where Group IIB metal cations, e.g., zinc cation, are incorporated into the zeolite. In this instance the platinum metal need not be introduced to obtain the selective hydroconversion, although its presence will add somewhat to the overall activity of the catalyst.

As an additional embodiment of the present invention, it has been found that the activity and effectiveness of the cation-exchanged zeolites hereinabove described can be substantially improved by contact with sulfur prior to their use in the treatment of the normal paraffin-containing hydrocarbon oils. The zeolite is preferably sulfactivated to enhance its paraffin removal properties by contact either with sulfur-containing feed or, if the feed has a low sulfur content, with hydrogen sulfide or an added sulfur compound which is readily convertible to hydrogen sulfide at the conditions employed, e.g., carbon disulfide and the like. The extent of this sulfactivation treatment should be sufficient to incorporate about 0.5 to 15 wt. percent sulfur into the zeolitic material. Where the catalyst does not contain an added hydrogenation component, but does contain the Group II-B metal, e.g., zinc suliactivation will be required.

For purposes of octane improvement, the feedstocks contemplated for use in the present process will generally be naphtha or high naphtha-containing feeds and may consist of either low boiling or high boiling naphthas. A typical low boiling feed has a boiling range of about 50 to 350 F., preferably 80 to 200 F., whereas the heavy naphtha has a boiling range of 257 to 550 F., preferably 300 to 450 F. These naphthas, both low boiling and high boiling, are exemplified by virgin naphtha fractions such as C -C naphtha, heavy virgin naphtha, heavy coker naphtha, heavy steam-cracked naphtha, heavy catalytic naphtha, etc.

For dewaxing purposes the feedstocks adapted for treatment in accordance with the present invention may be generally defined as hydrocarbon oils boiling in the range of about 300 to about 1100 F., and particularly between about 400 and about 650 J5. Such oils will include heavy naphthas, kerosenes (e.g., boiling between 300 and 500 F.), diesel fuels, jet fuels, heating oils, gas oils, middle distillates, lube base stocks, etc. The process of the invention is particularly effective for removing wax and similar normal paratfinic constituents from middle distillate and gas oil fractions, in order to reduce their pour point, cloud point, haze point and solidification tendency. The preferred middle distillate fractions will have a total n-parafiin content within the range of about 5 to 50 wt. percent, particularly 10 to 30 wt. percent; and the preferred gas oil fractions will have a total n-paraflin content within the range of about 10 to 50 wt. percent, particularly to wt. percent.

In accordance with the present invention, the feedstock is preferably preheated to the contacting temperature and introduced into contact with the crystalline zeolitic material, preferably in vapor phase or mixed vapor-liquid phase (in the case of a high boiling gas oil feed) and fixed bed operation, concurrently with a hydrogen-containing gas stream. Preferably the feed stream will pass downwardly through the bed of catalyst and, in so doing, the normal parafiins present therein are selectively hydrocracked to lower molecular weight products. The efiluent is removed from the zeolite catalyst contacting zone and passed to a vapor-liquid separation zone wherein the lower boiling normally gaseous hydrocarbons are removed and the normally liquid bottoms product is recovered and preferably further fractionated into light ends, naphtha and/or dewaxed fractions having desired boiling range, etc.

In carrying out the process of the invention, the operating conditions employed will depend upon the particular feedstock and end products. For octane improvement of naphtha feeds, typical conditions will include a temperature of 400 to 950 -F., preferably 650 to 850 F.; a pressure of 200 to 4000 p.s.i.g., preferably 500 to 2500 p.s.i.g.; a space velocity of 0.2 to 20, preferably 0.4 to 2, v./v./hr.; and a hydrogen rate of 1000 to 10,000, preferably 1500 to 5000, standard cubic feet of hydrogen per barrel of feed. For the gas oil or middle distillate feeds (dewaxing purposes), typical conditions will include a temperature of from about 650 to 900 F., preferably 700 to 850 F.; a pressure of 100 to 5000 p.s.i.g., preferably 400 to 1000 p.s.i.g.; a space velocity of 0.1 to 10, preferably 0.5 to 2, volumes of feed per volume of crystalline zeolite per hour; and in the presence of hydrogen which is preferably introduced concurrently with the feed at a rate of about 500 to 100,000, preferably 1000 to 3000, standard cubic feet per barrel of feed.

The following examples will serve to illustrate the advantages of the present invention and to set'forth the best mode now contemplated for carrying it out.

EXAMPLE 1 This example illustrates the preparation of the catalysts of the present invention. The starting material was a synthetic form of the mineral erionite, which was synthesized by known procedures and had elliptical pore openings of about 5.2 Angstroms on the major axis and a silica-toalumina mole ratio of about 7.

A 500 gram sample of the synthetic erionite was treated with an ammonium chloride solution, in order to exchange ammonium ions for the alkali metal cations (i.e., potassium and sodium cations) normally contained in the erionite as synthesized. The ammonium-exchange procedure was accomplished by suspending the sample in 1000 grams of water and adding 1954 grams of a 23 wt. percent ammonium chloride solution. The mixture was agitated for four hours, and the product removed by filtration and washed by suspension in 2000 grams of water with further agitation for about one hour. The above procedure was repeated three times so that the total number of exchanges was four.

Following the ammonium ion exchange, a portion of the erionite sample was further exchanged with a rare earth metal cation solution which consisted of an 8.7 wt. percent aqueous solution of didymium chloride. Two hundred grams of the DiCl -6H O salt were added to 300 grams of the ammonium form erionite (dry basis) suspended in 1800 grams of water. The slurry was heated to about F. and held for six hours with stirring. The hot slurry was filtered, washed with water and dried overnight at 250 to 350 F.

At this point 0.5 wt. percent palladium was incorporated into the rare earth metal-containing erionite product by suspending 126 grams of the oven dried material (112 grams dry basis) in 200 grams of water and adding 15 cc. of palladous ammonium chloride containing 37.5 mg. Pd/cc. thereto, with agitation. Agitation was continued for three hours at room temperature. The excess liquid was decanted and the material dried. The resultant catalyst showed the following analysis:

Wt. percent Alkali metal 2.4

Silica 74.0 Alumina 17.7 Palladium 0.5 Rare earth oxide 3.4

metal cationic exchange step was omitted. Analysis of this catalyst showed the following:

Wt. percent Alkali metal 2.5 Silica 76.7

Alumina 18.3

Palladium 0.5

Catalyst C was a similar relatively small pore size zeolite catalyst which was prepared from the natural mineral erionite by exchange with zinc cation. This catalyst was prepared as follows: a 316 gram sample of erionite (from a deposit in Pine Valley, Nev.) was suspended in 2000 grams of water. A solution of one pound zinc chloride in 500 grams of water was added thereto at room temperature, and the mixture was agitated for four hours. The product was removed by filtration and washed by suspension in 2000 grams of water with agitation for one hour. After filtration the above procedure was repeated twice, so that the total number of exchanges was three. The zinc-containing product was dried overnight at 250 to 300 F. and weighed 286 grams. Toincorporate palladium into the zinc-erionite product, the 286 grams of product were suspended in 600 grams of water and, with agitation, 50 cc. of palladous ammonium chloride containing 37.5 mg. Pd/oc. was added thereto. Agitation was continued for one hour at room temperature. After filtering, washing, drying and calcining, the catalyst analyzed 0.6 wt. percent palladium, 7.7 wt. percent zinc and had a SiO /Al O ratio of 7.3:1. Catalyst C represents a previously known superior selective hydrocracking catalyst.

EXAMPLE 2 Catalysts A, B and C of Example 1 were evaluated in the selective hydrocracking of a naphtha feed derived from an Arabian crude oil. The naphtha teed had a gravity of 853 API, a boiling range of 110 F. to 185 F., a normal pentane content of 17.0% and a normal hexane content of 35.1%. All catalysts were first sulfactivated by contact with the same feed containing 1.0 Wt. percent carbon disulfide. Fresh naphtha feed was then passed downwardly over a fixed bed of the sulfactivated catalyst in pellet form with the concurrent introduction of hydrogen. Conditions utilized included a space velocity of 0.5 volumes of feed pervolume of catalyst per hour, a pressure'of 500 p.s.i.g. and an exit hydrogen rate of about 2000 standard cubic feet per barrel. Catalyst performance was measured by the disappearance of normal parafiins by their conversion to C; gases. The results obtained with the three catalysts are summarized in the following Table I.

C represents a previously known highly superior seleo tive hydrocracking catalyst.

EXAMPLE 3 Catalyst A (palladium-didyrnium-hydrogen-erionite) of Examples 1 and 2 was used in the treatment of a Louisiana-Mississippi gas oil feed to obtain pour point reduction by selective hydroconversion of the waxy components. The waxy gas oil feed was passed downwardly through a fixed bed of the catalyst at a space velocity of 0.5 v./v./hr., an exit hydrogen rate of 2000 s.c.f./bbl., a pressure of 500 p.s.i.g. and at temperatures of 700 and 800 F. Cracked products boiling below 320 F. were removed from the effiuent. The yield of 320 F.+ product and its pour point and cloud point were determined. The data obtained are compared below to a conventional dewaxing operation using a mixed solvent of methyl ethyl ketone-methyl isobutyl ketone (in 3:1 ratio) in the ratio of 3 parts by volume solvent to 1 part gas oil feed. In these runs, the feed was heated to 150 F. and then chilled slowly while incrementally adding the solution mixture. At -5 F., the wax was filtered and dewaxed oil recovered. The data are summarized in the following Table II.

TABLE I.SELECTIVE HYD ROCRACKING OF ARABIAN C -C NAPHTHA [0.5 v./v./hr., 500 p.s.i.g., 2,000 s.c.t.[bbl. H2]

Catalyst A B 3 Palladium Content, wt. percent u 0.5 0.6

Predominating Cation Didym m-Hydrogen Hydrogen zin Temperature, F 85 750 700 150 700 800 750 700 Prodguct Analyses, wt. percent: Feeg I 3 7 56.6 48.8 72.9 61.5 63.8 56.7 17. 0 0.3 0. 5 1. 4 0. 9 2. 0 1. 9 4. 7 g n-C -1 -0 0.2 0.2 0.3 0.1 0 4 4 4 Conversion C5, percen 98 97 92 95 88 89 72 22 Conversion n-Ct, percent. 100 100 99 99 99 100 99 87 As indicated above, the catalyst of the invention, i.e., Catalyst A, exhibited superior activity and/or selectivity to the comparative Catalysts B and C. The above results may be evaluatedby comparison of (.1) the conversion of the nC and nC paraffins, (2) production of the C product and (3) the temperature required for near 100% conversion. Using these criteria, it is seen that Catalyst A is superior to Catalyst B in that higher conversion of normal paraflins was obtained r at the same temperature levels. Furthermore, not only is Catalyst A superior to Catalyst B in the higher conversion of normal paraffins, but it is to be observed that excessive cracking is minimized with Catalyst A. Thus, at the comparable levels of 750 and 700 F., the appearance of C product is approximately equal to' the normal parafiinic content of the feed in the case of Catalyst A; whereas in the case of Catalyst B, the C yield at these temperature levels is substantially higher, indicating excess undesired cracking of materials other than the objectionable normal paraffins. The superior activity and selectivity of the catalyst of the invention are thus demonstrated.

With respect to Catalyst C, Catalyst A is demonstrated to be markedly superior at comparable temperature levels, as evidenced, for example, by the n-C conversion figures. Furthermore, the results obtained with Catalyst A at a temperature of 700 F. are nonetheless superior to those obtained with Catalyst C" at 800 F., again attesting to the superiority of the catalyst of the invention. Thus, substantially lower temperatures can be utilized with catalysts of the invention, whereby excessive cracking of desired products can be minimized with maximum conversion of objectionable normal paraflinic constituents in the feed. It is again to be noted that Catalyst TABLE II.DEWAXING OF LA.-MISS. GAS OIL WITH PALLADIUM-RARE EARTH ERION'ITE [0.5 v./v./hr., 500 p.s.l.g., 2,000 s.c.f./bbl. H2]

The superiority of the catalysts of the present invention is demonstrated by higher yield, lower pour point and lower cloud point product for the 800 F. operation, as compared to standard solvent dewaxing.

EXAMPLE 4 Catalyst D was prepared by taking a 147 gram portion of oven dried didymium exchange-treated ammonium erionite prepared as described in Example 1 (equivalent to grams dry basis) and having a Di O content of 3.4 wt. percent, slurrying it in 1 liter of hot water and adding 46.3 grams of ZnSO- -7H O. The slurry was heated and stirred for 30 minutes, filtered, water Washed and oven dried. A portion of this material was analyzed and showed 2.2 wt. percent Di O and 13.8 wt. percent ZnO, indicating that some of the rare earth had been displaced by the zinc. One hundred thirty-two grams of the oven dried material (123 grams dry basis) were contacted with 220 cc. of a Pd(NH C1 solution 1, 1 containing 0.615 gramld. The slurry was stirredintermittently at'room temperature for threehours .and then oven dried. The catalyst contained 0.5 wt. percent Pd. It was formed into A x cylindrical pellets which were cracked, broken and sized l4 65 mesh.

EXAMPLE 5 Temp, F.

Feed Selective Solvent Hydrocracklng Dewaxing 320 F.+ Yield, Wt. percent- 100 88.9 '83. 2 80 Inspections on 320 F.+ Fraction:

Gravity, API at 60 F. 33.5 32. 7 31. 6 31. 9 Pour Point, F 40 45 -10 Cloud Point, F 40 18 0 The superiority of the hydroselective cracking with the catalysts of the present invention is again demonstrated by higher yield and lower pour point product for the 800 F. operation as compared to standard solvent dc waxing operation carried out at -5 F.

What is claimed is: 1. A process for selectively hydrocracking straightchain hydrocarbons contained in a hydrocarbon feedstock which comprises contacting said feedstock at hydrocracking conditions in the presence of hydrogen with a catalyst comprising a rare earth metal-containing crystalline alumino-silicate zeolite of the erionite variety having uniform pore openings less than about 6 Angstrom units.

2. The process of claim 1, wherein said feedstock is selected from the group consisting of naphtha fractions, gas oil fractions and middle distillate fractions.

3. The process of claim 1, wherein said zeolite is synthetic erionite and said catalyst additionally comprises a metallic hydrogenation component selected from the group consisting of metals in Groups V-B, VI-B, VII-B and VIII of the Periodic Table.

4. The process of claim 3, wherein said metallic hydrogenation component comprises a platinum group metal.

5. The process of claim 1, wherein said zeolite additionally contains a Group II-B metal. 7

6. The process of claim 5, wherein said Group II-B metal is zinc.

7. The process of claim 1, wherein said zeolite has been base exchanged with hydrogen-containing cations.

8. The process of claim 3, which additionally comprises sulfactivating said catalyst by contact with a sulfur compound.

9. The process of claim 5, which additionally comprises sulfactivating said catalyst by contact with a sulfur compound.

10. A process for improving theoctane rating of naphtha'fractio'nsby selectively hydrocracking straight-chain hydrocarbons contained in said naphtha fractions which comprises contacting said naphtha fractions at hydrocrackin'g conditions infthe presence (or hydrogen with a catalyst comprising a metallic hydrogenation component combined with a "rarc' e'arthmetal containing crystalline alumino-silicate zeolite of the erionite variety having uniform pore openings of about '5 Angstrom units and recovering a naphtha product of improved octane rating.

11. The processof claim 10, wherein said metallic hydrogenation component is a platinum group metal.

12. The process of claim-10,-whichadditionally comprises contacting said catalyst withra sulfur-containing compound. a

13. The process of claim 10,.wherein said zeolite additionally-contains cations selected from-the group-consisting of Group II-B-metal cations, hydrogen-containing cations andmixturesthereof.

1 '14. A process for dewaxing petroleum oil fractions containing straight-chain hydrocarbons. which comprises contacting said fractions at conversion conditions in the presence of hydrogen with a crystalline alumino-silicate zeolite of the erionite variety having uniform pore openings of about 4 to less than 6 Angstrom units, said zeolite being combined with a metallic hydrogenation component and further containing a rare earth metal, and recovering dewaxed normally liquid product of substantially reduced pour point. V

15 The process of claim 14, wherein said metallic hydrogenation component is a platinum group metal. 7

16. The process of claim 14, which additionally com prises contacting said catalyst with a sulfur-containing compound.

17. The process of claim 14, wherein said zeolite additionally contains cations selected from the group consisting of Group I I-V metal cations, hydrogen-containing cations and mixtures thereof.

" 18. A catalyst composition comprising a metallic hydrogenation component combined with a rare earth metalcontaining crystalline alumino-silicate zeolite of the erionite variety having uniform pore openings of about 4 to less than 6 Angstrom units.

'19. The composition of claim 18 wherein said zeolite is synthetic erionite and said metallic hydrogenation component comprises a Group VIII metal.

20. The composition of claim 18, wherein said zeolite additionally contains cations selected from the group consisting of GroupII-B metal cations, hydrogen-containing cations and mixtures thereof.

References Cited UNITED STATES PATENTS 3,039,953 6/1962 A Eng 208-26 3,114,696 12/1963 WeisZ 208 -66 3,210,267 10/ 1965 Plank et al. 208- 3,240,697 3/1966 Miale ct a1. 208-120 3,268,436 8/1966. 'Arey et a1. 208--l1l DELBERT E. GANTZ, Primary Examiner.

ABRAHAM RIMENS, Assistant Examiner. 

